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SIMULATION OF METHYLDIETHANOLAMINE-
CARBON DIOXIDE-WATER SYSTEM USING
EQUILIBRIUM APPROACH
A THESIS SUBMITTED IN PARTIAL FULFILLMENT OF THE
REQUIREMENTS FOR THE DEGREE OF
Bachelor of Technology
in
Chemical Engineering
Submitted by
Vishal Mohanty
110CH0394
Under the guidance of
Dr. MADHUSHREE KUNDU
Department of Chemical Engineering
National Institute of Technology
Rourkela
2014
i
National Institute of Technology, Rourkela
CERTIFICATE
This is to certify that the thesis entitled, “SIMULATION OF
METHYLDIETHANOLAMINE- CARBON DIOXIDE -WATER SYSTEM USING
EQUILIBRIUM APPROACH”, submitted by Mr. Vishal Mohanty, Roll no. 110CH0394, in
partial fulfillment of the requirements for the award of degree of Bachelor of Technology in
Chemical Engineering at National Institute of Technology, Rourkela is an authentic work carried
out by her under my supervision and guidance.
To the best of my knowledge, the matter embodied in the report has not been submitted to any
other University / Institute for the award of any Degree or Diploma.
Date: Dr. Madhushree Kundu
Place: Rourkela Department of Chemical Engineering
National Institute of Technology
Rourkela – 769008
ii
ACKNOWLEDGEMENT
I express my sincere gratitude to Dr. M.Kundu for providing me an opportunity to work on this
project and for her constant guidance and timely suggestions throughout. I am also thankful to
Prof. H.M.Jena and Prof. R.K.Singh (Project Co-ordinators) for their valuable guidance and
advice.
Finally I thank my parents, sister and friends for their support and encouragement without which
this project would not have been possible.
Date: Vishal Mohanty
110CH0394
Department of Chemical Engineering
National Institute of Technology, Rourkela
iii
ABSTRACT
As the climate is changing very frequently and global warming is increasing at an alarming rate,
there is a need to mitigate greenhouse gases which have gained a top most priority. As the main
source of CO2 emission is fossil fuels (mainly in power plants), there is a need to keep check on
utilization of fossil fuels. Carbon capture and sequestration (CCS) is one of the methods which
enable the utilization of fossil fuels with lower CO2 emissions. Commercially CO2 capture by
chemical absorption is very active. While number of solvents for CO2 absorption by chemical
method has been proposed, comparison on performance of different solvents has not been done
properly and claims on different solvent vary widely. The work done here emphasize on
absorption of CO2 by the solvent and then stripping the solvent to obtain pure CO2.This thesis
has been prepared to evaluate the performance of solvent – methyldiethanolamine(MDEA) on
CO2 absorption.
In this thesis, comprehensive flow sheet model has been built for the solvent system, using
ASPEN Plus as the modeling tool. The Thermodynamic model using ENRTL model for CO2
capture has been done and effect of reboiler duty, solvent flow rate and so on has been observed.
iv
CONTENTS
Chapter Topic Page
no.
Abstract iii
List of figures vi
List of tables vii
Nomenclature viii
Chapter-1 Introduction 1-6
1.1 Carbon capture and sequestration 1
1.2 Brief overview of CO2 capture systems 2
1.2.1 Post combustion capture 2
1.2.1.1 Chemical absorption 3
1.2.1.2 Physical absorption 3
1.2.1.3 Membrane separation 3
1.2.2 Oxyfuel combustion 4
1.2.3 Precombustion capture 5
1.3 Thesis objective 5
1.4 Scope of the thesis 5
1.5 Outline of the thesis 5
Chapter -2 Literature review 7-10
2.1 Amine Process 7
2.2 Amine Type 7
2.3 Reaction Mechanism 8
Chapter-3 Methyldiethanolamine system 11-17
3.1 Detailed description of the process 11
3.1.1 Cooling of flue gas and compression 11
3.1.2 CO2 absorption and solvent regeneration 12
3.1.3 Solvent reclaimation 13
3.1.4 CO2 compression 13
3.2 Modeling used 13
v
3.2.1 Thermodynamic model 13
3.2.1.1 Solution chemistry and equilibrium governing equations 14
3.2.1.2 Electrolyte-NRTL model 16
3.2.2 Rate model 17
Chapter-4 Simulation By Thermodynamic Modeling 18-22
4.1 Procedure For Developing The Flowsheet 19
Chapter-5 Results and Discussions 23
5.1 Variation of L/G in Absorber with lean loading for 80% capture 23
5.2 Variation of Reboiler Duty with L/G in absorber for 80% capture 24
5.3 Effect of Capture Percentage 25
5.4 Variation Of Reboiler Duty with Pressure and Temperature of
Desorber for 80% capture
26
5.5 Variation of reboiler Duty with Lean Loading for Approach
Temperatures of 5℃ and 10℃
28
Chapter-6 Conclusion and Future Scope of Work 30
6.1 Conclusion 30
6.2 Future Scope of Work 30
References 31-33
vi
LIST OF FIGURES
Fig no.
Name of the figure
Page no.
1-1
Plot of global instrumental temperature anomaly vs.
Time
1
1-2
Plot of atmospheric CO2 concentration (ppmv) vs.
time as measured at Mauna Loa, Hawaii. Data from
2
1-3
Schematic of post-combustion capture
3
1-4
Schematic of oxyfuel combustion
4
1-5
Schematic of precombustion decarbonization
5
2-1
Reaction of tertiary amines with CO2
9
2-2
Zwitterion mechanism
9
2-3
Alkyl carbonate formation mechanism
9
3-1
Schematic of CO2 capture by use of MDEA solvent
11
4-1
Process flow diagram of MDEA system as
developed in ASPEN Plus
18
4-2
Different components selected in ASPEN PLUS
19
5-1
Variation of L/G with lean loading for 80% CO2
capture
23
5-2
Variation of reboiler duty with L/G for 80% CO2
capture
24
5-3
Variation of reboiler Duty with Lean Loading for
different Capture percentage
26
5-4
Graph showing variation of reboiler Duty with
desorber Pressure
27
5-5
Graph showing variation of desorber temperature
with desorber Pressure
27
5-6
Graph showing variation of reboiler Duty with
cross heat exchanger approach
28
vii
LIST OF TABLES
Table no.
Name of the table
Page no.
1-1
CO2 partial pressure in flue gases of different combustion
systems
4
4-1
Input Specification Of Pumps
20
4-2
Input Specification of Heaters
20
4-3
Reaction Data Used for the Reactions
21
5-1
Lean Loading vs L/G Data
23
5-2
Data representing variation of Reboiler Duty with L/G in
Absorber
24
5-3
Variation of reboiler duty with Lean loading at different
capture percentage
25
5-4
Data Representing variation of Reboiler duty with
temperature and pressure of the desorber for 80% capture
26
5-5
Data representing Variation of Reboiler duty with Lean
Loading for different temperature approach
28
viii
NOMENCLATURE
r : Rate of a reaction
k2 : Second-order rate constant
R : Universal gas constant
T : Temperature
Kj : Equilibrium constant for the reaction j
ΔGj°(T) : Gibbs free energy change for reaction j at temperature T
s : Solvent
Gs(T) : Solvent s Gibbs free energy at temperature T
Gsig
(T) : Solvent s ideal gas Gibbs free energy at temperature T
ΔGsigfl
(T) : Departure of Gibbs free energy from ideal gas to liquid at temperature T
ΔHabs : Heat of absorption/mole of CO2
Hl
final : Final solution’s molar enthalpy
Hlinitial : Initial solution’s molar enthalpy
Hg CO2 : Gaseous CO2 absorbed molar enthalpy
nfinal : Total number of moles in the final solution
ninitial : Total number of moles in the initial solution
¥ : Activity coefficient of species i in solution
ni : Number of moles of i
GE : Excess Gibbs free energy
P : Pressure
Gid
: Excess Gibbs free energy if the mixture were ideal
L : Total molar flow of liquid
G : Total molar flow of gas
1
CHAPTER-1
INTRODUCTION
1.1 CARBON CAPTURE AND SEQUESTRATION
As the climate change fear has achieved a great importance, various methods have been
developed to mitigate CO2 emission. For the last 5 decades there has been growing concern as
the average global temperature is increasing at a moderate rate. Figure 1-1 tells us about the rise
in the difference between the global mean surface temperature and the average temperature from
1961-1990 [1].
Figure 1-1: Plot of global instrumental temperature anomaly vs. time (temperature average
from 1961-1990) [1].
The main cause behind increase in CO2 is due to various activities adopted by human beings
which directly or indirectly contribute to the rise in CO2 concentration.
From figure 1-2 it can be seen that CO2 concentration is increasing at a concerned rate. So there
is need to mitigate CO2 emissions.
2
Plot of atmospheric CO2 concentration (ppmv) vs time as measured at
Mauna Loa, Hawaii
390
Co
nce
ntr
ati
on 380
370
(pp
mv
)
360
2
350
Atm
osp
her
ic C
O
340
330
320
310
1958 1963 1968 1973 1978 1983 1988 1993 1998 2003 2008
Time
Figure 1-2: Plot of atmospheric CO2 concentration (ppmv) vs. time as measured at Mauna
Loa, Hawaii [2].
The main source for the increase in CO2 concentration in the atmosphere is electricity generation
sector which mainly comprises of Natural gas and Coal fired power plants. So there is a urgent
need of deploying CO2 mitigation technologies on this sector which will provide a cleaner
environment in future.
1.2 BRIEF OVERVIEW OF CO2 CAPTURE SYSTEMS
Three processes for CO2 capture have been classified which are as follows:-
i)Post-Combustion Capture
ii)Oxyfuel Combustion
iii)Pre-Combustion Capture
1.2.1 POST-COMBUSTION CAPTURE
In this process, CO2 is removed from the flue gas produced after combustion of the fuel. A
schematic of post-combustion capture is presented in Figure 1-3
3
Figure 1-3: Mechanism of post-combustion capture [3]
Various post-combustion methods are Chemical absorption, Physical absorptions, Membrane
separation, Adsorption, Cryogenic separation.
1.2.1.1 CHEMICAL ABSORPTION
CO2 is removed from the flue gas by passing the gas through the continuous scrubbing system.
Usually, an absorber and a desorber together make the scrubbing system. In the absorber,
reversible reaction between CO2 and solvent (amine) takes place. In the desorber, the CO2 which
was absorbed by the solvent (amine) is stripped off and a pure stream of CO2 is produced, which
is then sent for compression and the regenerated solvent is sent back to the absorber.
1.2.1.2 PHYSICAL ABSORPTION
Physical absorption which is also known as non-reactive absorption takes place between two
phases, either the gas is absorbed by the liquid, or the liquid is absorbed by a solid.
When the gas is absorbed by the solvent in the liquid phase, some amount of gas moves into the
liquid. For example, oxygen may be absorbed by the water. At the liquid gas interface the mass
transfer is taken place and the rate is dependent both upon the gas and the liquid. The absorption
is dependent on gas solubility, the pressure and the temperature. When the liquid is absorbed by
a solid, some amount of liquid moves into the solid. For example, water may be absorbed by the
pot (clay) in which it is stored. At the liquid solid interface the mass transfer is taken place and
the rate is dependent both on the liquid and the solid.
1.2.1.3 MEMBRANE SEPARATION
The process by which materials are selectively separated through pores or minute gaps in the
molecular arrangement of a continuous structure is known as Membrane separation. Membrane
separations are classified on the basis pore size as well as the separation driving force.
4
Table 1-1: CO2 partial pressure in flue gases of different combustion systems. [3]
Flue gas source
CO2 Pressure of gas CO2 partial
concentration, stream, pressure,
% vol (dry) MPa MPa
Natural gas fired 7-10 0.1 0.007-0.01
Boilers
Gas turbines 3-4 0.1 0.003-0.004
Oil fired boilers 11-13 0.1 0.011-0.013
Coal fired boilers 12-14 0.1 0.012-0.014
IGCC after 12-14 0.1 0.012-0.014
combustion
IGCC synthesis gas 8-20 2-7 0.16-1.4 (before shift)
after gasification
1.2.2 OXYFUEL COMBUSTION
Here instead of air, combustion is taken place in the presence of oxygen which prevents dilution
of CO2 with other gases.
Figure 1-4: Mechanism of oxyfuel combustion [3]
5
1.2.3 PRE-COMBUSTION CAPTURE
In this process, carbon content of the fuel is reduced before combustion so as to produce pure
CO2 on combustion. Figure 1-5 presents a schematic of precombustion decarbonization.
Figure 1-5: Mechanism of precombustion decarbonization [3]
1.3 THESIS OBJECTIVE
The main objective of the thesis is to develop an energy efficient absorption and capture of CO2
using methyldiethanolamine (MDEA).
1.4 SCOPE OF THE THESIS
The scopes of this thesis are as follows:
1. To develop a thermodynamic model which can describe thermodynamic and thermal
properties of CO2-H2O-alkanolamine system over different pressure, temperature and
amine concentration range
2. To determine the parameters which affect the performance of the CO2-H2O-alkanolamine
system
3. To optimize the different dependent parameters to get the best results and maximum CO2
capture percentage
1.5 OUTLINE OF THE THESIS
The work is divided into 7 chapters. Chapter 1 of this thesis tells about the carbon capture and
sequestration and various process involved with it and about the thesis objectives. Chapter 2
involves the studies of the previous work that has been done on the given topic and discussion
about the amines involved and the reaction mechanism. Chapter 3 describes about the
methyldiethanolamine system. This chapter also tells about the actual process going on and the
6
model used to describe the system. Chapter 4 shows the developed thermodynamic model for
CO2-H2O-MDEA and tells about the various parameters (inputs) used for modeling the system.
Chapter 5 describes the result and discussion for this work. It tells about the effect of various
parameters on the system. Chapter 6 presents overall conclusions from this thesis. Chapter 7
discusses further direction in which this study can be carried forth.
7
CHAPTER-2
LITERATURE REVIEW
By reviewing work from previous studies, it provides a motivation and direction for this work.
Additionally, some of the results will be comparable to the results seen later in this work.
Upon surveying it was found that more than 60 years ago in chemical industries (mostly oil
industries, power plant) Amine scrubbing technology for removal of H2S and CO2 from various
gas streams was adopted. This technology commercially is the most widely used process for
capturing CO2.
2.1 AMINE PROCESS
In most of the industries absorption with the help of chemical solvents, which is also known as
chemical absorption is the commercially most widely used process to remove acid gas(mainly
CO2 and H2S) from various gas streams. Currently preferred chemical solvents for acid gas
removal by chemical absorption are amine based absorbents.
Alkanolamines, which are the combinations of alcohols and ammonia, are the mostly preferred
solvents for removing acid gas. In addition to natural gas processing, chemical absorption of acid
gases by alkanolamines has been utilized in a various industries like petroleum refining, CO2
capture from combustion and flue gases, removal of CO2 from synthesis gas in ammonia or
hydrogen plants.
2.2 AMINE TYPE
Though there are many amines available but commercially 4 types of amines are used which are
primary amines (e.g. MEA), secondary amines (e.g. DEA), tertiary amines (e.g. MDEA) and
cyclic amines (e.g. PZ) [4].
As there are three alkyl groups present in MDEA, the reaction between CO2 and MDEA is
hindered. Therefore, CO2 dissolves in H2O first to form carbonic acid (H2CO3), and then
carbonic acid reacts with amine. The following reaction takes place
CO2 + H2O + MDEA MDEAH+ + HCO3
- -- (2.1)
For determining the solubility of CO2 in MDEA as well as in other Alkanolamines, the ENRTL
(electrolyte NRTL) [5-7] was used by Posey [8] and Austgen et al. [9]. For determining the VLE
data for the MDEA- H2O-CO2 system, The Pitzer’s equation [10] was used by Kuranov et al.
[11] Kamps et al. [12] and Ermatchkov et al. [13]. For determining the Solution Enthalpy of CO2
in aq. MDEA, the data for VLE were combined with the Pitzer’s equation by Arcis et al. [14]
8
and the thermodynamic model was applied. For analyzing the VLE for absorption of CO2 in aq.
MDEA, The extended UNIQUAC model [15] was used by Faramarzi et al. [16]. Concentration
of MDEA solutions containing CO2 was also predicted and was compared to the concentration
measured by NMR spectroscopy [17, 18].
Nowadays aqueous solutions of MEA are used mainly. The major advantages of MEA are: large
reactivity, cheap, and less capacity for absorption of hydrocarbons. The main disadvantages of
MEA are: High corrosiveness of MEA which increases by temperature, large heat of reaction
with CO2 and H2S results in large energy requirements for the regeneration of solvent and
subsequently the total cost for the process is raised, relatively large vapor pressure which sees
amine losses through vaporization [4].
Though MDEA is costlier than MEA, but it has lower rate of reaction with CO2 as compared to
other amines which makes it suitable to use in various industries.
The main reactions taking place in CO2-MDEA-H2O system are as follows [4]:
Water dissociation equilibrium:
2H2O H3O+ + OH
- -- (2.2)
Bicarbonate formation equilibrium:
CO2 + 2H2O H3O+ + HCO3
- -- (2.3)
Carbonate formation equilibrium:
H2O + HCO3- H3O
+ + CO3
2- -- (2.4)
MDEA protonation equilibrium:
H3O+ + MDEA H2O + MDEA
+ -- (2.5)
2.3 REACTION MECHANISM
The rate equation for the reaction of tertiary amines (MDEA) with CO2 is given by the equation
[19]
r=k2[MDEA][CO2] -- (2.6)
9
Figure 2-1: Reaction of tertiary amines with CO2
Direct reaction of CO2 with amine is not possible. Therefore, water must be present so that the
reaction takes place by above mechanism [19].
Barth et al. proposed two other mechanisms for the reaction of amines with CO2 in addition to
simple second order mechanism. The first is the zwitterion mechanism, in which there is a
possibility of forming an intermediate [19].
Figure 2-2: Zwitterion mechanism
The second mechanism is the formation of alkyl carbonates which has got a rare chance to occur
in solutions with high pH.
Figure 2-3: Alkyl carbonates formation mechansim
The reason that MDEA acts well is that it does not react with CO2 appreciably. Therefore, it
absorbs this gas very slowly. Certain additives like piperazine are used so as to increase the rate
10
of reaction between CO2 with MDEA. By adding required amount of additives to MDEA,
desired amount of CO2 can be removed.
Posey analyzed that Absorption of CO2 into the aqueous MDEA solutions is controlled by mass
transfer mechanism with chemical reaction. Therefore, Solubility is required for calculating the
mass transfer driving force. Thus for predicting equilibrium solubility at desired conditions of
temperature, MDEA concentration and CO2 loading, a thermodynamic equilibrium is highly
demanded. Bishnoi also suggested that for calculating free amine concentration, thermodynamics
is also needed. To determine the quantity of free amine for a specified overall acid gas
concentration Chemical equilibrium calculations for all the species present in the liquid phase is
needed. More is the amount of available free amine, faster is the rate of reaction between amine
and acid gases [4].
Blauwhoff et al. studied MDEA using stirred tank. He found a 2nd
order rate constant of
4.8m3/kmole/sec at 293K.he reported that it is hard to get a constant absorption rate with MDEA.
He assumed 1st order mechanism with respect to MDEA. Critchfield and Rochelle showed
results of CO2 absorption over a range of temperature (282-350K). The nominal pressure for CO2
was 1 atm [19].
Toman and Rochelle studied the absorption of CO2 into 50% MDEA and another solution with
same strength neutralized with sulfuric acid. The rate constant was to found to be little higher
which suggested a weak catalytic effect. Hence, Glasscock proposed a rate expression that
showed this decrease as a consequence of depletion of hydroxide ion at the interface [19].
r=[CO2][MDEA]{KH2O(H2O) + KOH-(OH
-)} + K
’OH
-[CO2][OH
-] -- (2.7)
11
CHAPTER-3
METHYLDIETHANOLAMINE SYSTEM
3.1 DETAILED DESCRIPTION OF THE PROCESS
The process used for capturing CO2 using MDEA can be divided into 3 different sections:
1. Cooling of Flue gas and its compression
2. Absorption of CO2 and solvent regeneration
3. Compression of CO2
A detailed flowsheet of the process is given below:
Figure 3-1: Flowsheet for CO2 capture by MDEA solvent [20]
3.1.1 COOLING OF FLUE GAS AND COMPRESSION
The absorber for the CO2-H2O-MDEA system should operate at temperature of around 40°C and
therefore, the gases temperature at the inlet of the absorber should lie in the temperature ranging
from 40-50°C. Usually, the flue gases temperature at the exhaust in industries ranges from 110-
120°C and hence, the flue gases must be cooled before feeding it to the absorber. Sometimes wet
flue gas desulfurization scrubber is used for cooling the flue gases. If the flue gases have not
been through a scrubber then cooling is done by other means.
12
Direct contact cooling tower (DCC) is used for cooling the flue gases by feeding it the
tower.DCC may be tray tower or packed tower where counter-current flow of flue gases and the
cooling water takes place. The flue gas is fed from the bottom of the tower whereas cooling
water is fed from the top. Inside the tower, the flue gas is cooled by water evaporation and hence,
at the exit the water content of the gas is diminished at the exit of the tower. From the bottom of
the tower the cooling water is collected and is sent again to another cooling tower for lowering
the temperature so that it can be used gain in the DCC.
The flue gas coming out of the DCC needs to be compressed and therefore, it is sent to a blower.
Because of the upward movement of the flue gas in the absorber (tray column), the pressure of
the flue gas needs to be raised before feeding to the absorber. Along with the pressure,
temperature is also increased. Flue gas needs to be scrubbed prior to chemical absorption with
MDEA to remove NOx, SOx and other impurities, which react irreversibly with MDEA to form
heat stable salts which cannot be reclaimed. The recommended concentration for NO2 should be
less than 20 ppmv [21]. Similarly the recommended concentration for SOx should be less than 10
ppmv for MDEA solvent [21]. A wet electrostatic precipitator or a mist eliminator must be
employed in the flue gas desulfurization unit [21] so as to remove SO3, which can form sulfuric
acid aerosol in scrubbers which can cause corrosion.
3.1.2 CO2 ABSORPTION AND SOLVENT REGENRATION
The absorber used is a tray column, where vapor and liquid leaving the stage are in equilibrium.
From the absorber’s bottom, flue gas is fed whereas lean amine solvent is fed from the top. The
loading of the lean amine stream which is entering to the absorber from the top is between 0.3-
0.35 and rich amine stream leaving the absorber has a loading close to 0.8. In a MDEA system,
the loading is defined on a mole basis as given by
Moles of CO2
Loading = -- (3.1)
Moles of MDEA
The amine stream which is stripped off CO2 is referred to as Lean amine off i.e. the amine stream
entering from the absorber’s top. If the amine stream has CO2 loaded in it, then it is known as
rich amine i.e. the stream leaving from the absorber’s bottom. The lean amine stream is entering
inside the absorber through the second stage from the top and make-up water is entering at the
top stage.
The rich amine leaving at the absorber’s bottom is sent to the heat exchanger, which is also
known as cross-heat exchanger, via a pump. In the cross heat-exchanger, the heat is exchanged
between the rich amine stream from the absorber and the lean amine stream from the desorber.
13
As a result, rich stream gets heated up and lean stream gets cooled down. Before the lean amine
stream is fed to the absorber, it is again cooled to bring the temperature below 40°C.
The desorber is fed with the rich stream from the cross-heat exchanger. The desorber used here is
a tray column which has a reboiler of kettle type. The desorber typically operates at slightly
elevated pressure of 1.7-1.8 atm. The rich amine is fed at the second stage of the desorber and
flows downward inside the column, opposite to the vapors’ direction from the reboiler. The
stream that is coming out from the top of the desorber is sent to a condenser where water is
condensed and the temperature is lowered and then sent to a flash to separate the CO2 from H2O
[22].
3.1.3 SOLVENT RECLAIMATION
A MDEA reclaimer is used necessarily when the flue gas is coming out from coal fired power
plants to treat the stable salts that are formed because of SOx and NOx. These salts should not
accumulate on the amine stream as it is not desirable since it reduces the solvent’s capacity for
absorbing CO2. The removal of the purge stream of MDEA solvent is done and is then sent to the
reclaimer where strong alkali like NaOH and heat are added because of which heat stable salts
can be dissociated contributing to the recovery of the solvent [22].
3.1.4 CO2 COMPRESSION
The CO2 gas coming out from the desorber’s top needs to be dried and should be compressed
before sending it for storage. Drying is one of the important steps as the presence of even small
amount moisture in the stream can corrode the pipelines, which are used for transporting CO2.
Typically, a reciprocating compressor with 4 stages is used for cooling. The compressor is
employed for compressing the CO2 to a pressure of 90 atm, after which the liquid CO2 can be
pumped through pump with the discharge pressure of 130 atm [22].
3.2 MODELLING USED
Types of modeling used
1. Thermodynamics modeling
2. Rate based modeling
3.2.1 THERMODYNAMIC MODEL
The thermodynamic model was described along with the help of electrolyte-NRTL frame work.
A model for the Absorption/Stripping process demands the use of rigorous thermodynamics.
Even though CO2 absorption is a non-linear process, information about the equilibrium
composition of the CO2-MDEA-H2O system is required for determining driving forces in liquid
phase and solution speciation [23].
14
3.2.1.1 SOLUTION CHEMISTRY AND EQULIBRIUM GOVERNING EQUATIONS
The following reversible reactions occur in the liquid phase when CO2 is absorbed into an
aqueous solution of MDEA [24]
R3NH+ = R3N + H
+ -- (3.2)
CO2 + H2O = H+ + HCO3
- -- (3.3)
HCO3- = H
+ + CO3
2- -- (3.4)
H2O = H+ + OH
- -- (3.5)
Corresponding expressions of equilibrium rate constants are:
K1= ( aR3N aH+)/(aR3NH) -- (3.6)
K2 = ( aH+ aHCO3-)/( aCO2 aw ) -- (3.7)
K3 = ( aH+ aCO32-
)/( aHCO3-) -- (3.8)
K4 = ( aH+ aOH-)/(aw) -- (3.9)
The equilibrium constants of the above mentioned reactions are calculated from the reference-
state participating components’ Gibbs free energies:
-RT ln Kj = ΔGj(T) -- (3.10)
Where Kj denotes the equilibrium constant for the reaction j, ΔGj°(T) denotes the Gibbs free
energy change for reaction j reference-state at temperature T, R denotes the universal gas
constant, and T denotes the temperature of the system [25].
For the aqueous phase reactions, the reference states chosen for the solvents (water and MDEA)
are pure liquid, and for the solutes at infinite dilution. The calculation of the solvents’ Gibbs free
energy is done by the function shown below:
Gs(T) = Gsig
(T) + ΔGsigfl
(T) -- (3.11)
where Gs(T) denotes the solvent s Gibbs free energy at temperature T, Gsig
(T) denotes the
solvent s ideal gas Gibbs free energy at temperature T, and ΔGsigfl
(T) denotes the departure of
Gibbs free energy from ideal gas to liquid at temperature T.
The calculation of the Gibbs free energy for an ideal gas is done from the Gibbs free energy of
formation for an ideal gas at 298 K, the enthalpy of formation for an ideal gas at 298 K, and heat
capacity for the ideal gas [25].
15
The heat of absorption for CO2 in aqueous MDEA solutions can be done from the enthalpy
balance for the absorption process:
ΔHabs = nfinalHlfinal – ninitial H
linitial – nCO2H
gCO2 -- (3.12)
nCO2
where ΔHabs denotes the heat of absorption/mole of CO2, Hl
final denotes the final solution’s molar
enthalpy, Hlinitial denotes the initial solution’s molar enthalpy, H
g CO2 denotes the gaseous CO2
absorbed molar enthalpy, nfinal denotes the total number of moles in the final solution, ninitial
denotes the total number of moles in the initial solution, and nCO2 denotes the number of moles of
CO2 absorbed [25].
There are basically two types of heat of absorption which are 1) integral and 2) differential heat
of absorption. The integral heat of absorption for the amine-H2O-CO2 system is the heat
effect/mole of CO2 during the CO2 loading of the amine increasing from 0 to the final CO2
loading value of the amine-H2O-CO2 system. The differential heat of absorption for the amine-
H2O-CO2 system is the heat effect/mole of CO2 if meager amount of CO2 is added to this amine-
H2O-CO2 system.
For calculating both the types of heat of absorption, enthalpy calculations for the initial and final
amine-H2O-CO2 systems and for gaseous CO2 are needed. The calculation for heat capacity of
the MDEA-H2O-CO2 system can be done from the temperature derivative of enthalpy [25]
The following balance equations for the reacting species can be formed [24]:
Amine balance:
MA = (mR3N + mR3NH+) -- (3.13)
Carbon balance:
MAy = (mCO2 + mHCO3- + mCO3
2-) -- (3.14)
Electroneutrality:
mR3NH+ + mH+ = mOH- + mHCO3
- + 2mCO3
2- -- (3.15)
16
3.2.1.2 ELECTROLYTE-NRTL MODEL
The thermodynamic model developed in this work uses the electrolyte-NRTL theory which was
first developed by Chen and Evans in the year 1979 [26, 27] and extended by Mock et al [28, 29]
in the year 1986 for mixed solvent electrolyte systems.
The E-NRTL model is a used to determine the excess Gibbs free energy for a solution. It is
found that the activity coefficients of every component of mixture can be related to the excess
Gibbs free energy by equation [23]
Ln¥i=d/dni(GE/RT)T,P,n j≠I -- (3.16)
¥ is the activity coefficient of species i in solution, ni is the number of moles of i, GE
is the excess
Gibbs free energy, defined as
GE=G-G
id -- (3.17)
Where Gid
is the excess Gibbs free energy if the mixture were ideal.The excess Gibbs free energy
is related to excess enthalpy and excess entropy of mixing by
GE=H
E-TS
E -- (3.18)
The excess enthalpy is found from the fact that a species changes its interactions with the
surrounding species when the composition is changed. When ions are present in larger amounts
in the solution, the interaction with each other is strong and every molecule of CO2 among them
diminishes the intensity of the interaction. It tends to leave the liquid phase when the solution has
high ionic strength because this reduces the total enthalpy of the solution. The excess entropy is
due to the change in the randomness of the reciprocal position of molecules in the solution. The
electrolyte-NRTL model assumed the non ideal entropy of mixing is very less as compared to
heat of mixing. It gives an analytical expression for the excess Gibbs free energy. In the model
GE is a sum of three terms, the long range ionic forces, the short range molecular forces and the
Born correction [24].
GE=G
E, PDH+G
E, BORN+G
E, NRTL -- (3.19)
The long term ionic is described with the theory of Debye-Huckel, modified by Pitzer and it
depends on ionic strength of the solution. The term GE, PDH
+GE, BORN
represents the long range
forces contribution with reference state of ions at infinite dilution in water.
The short range force needs to be included for accounting the hydrogen bonds as well as local
interactions of molecules with molecules, molecules within parts and ion pairs with ion pairs
[24].
17
3.2.2 RATE MODEL
The modeling of absorption and desorption column is done by the use of thermodynamic models.
In the thermodynamic models, the column is defined with specific number of stages and it is
assumed that the liquid and vapor leaving each stage are in equilibrium [30]. Though, this is an
important assumption, it may not be valid in real case scenario. However to make it applicable
for real case scenarios, factors like the height equivalent to a theoretical plate (HETP), Murphee
and stage efficiencies are used [31]. Even after the above factors are used, it does not work
properly as there are large deviations from the equilibrium model [32] .Therefore, t is necessary
to use the rate-based models.
The modeling of absorption and desorption columns by rate-based model is done by the rate-
based mode of RADFrac, ASPEN RateSep. In this type of modeling chemical reactions kinetics
as well as heat and mass transfer phenomena are included [33]. By the help of Aspen RateSep,
various equations are solved, which are:
• Balance equation for Heat and mass both for the liquid and vapor phases.
• Rate models for Heat and mass transfer to determine the transfer rates at the
interphase.
• Equations for Vapor-liquid equilibrium at the interphase.
• Estimation of coefficients for heat and mass transfer as well as interfacial areas.
• Improvement of heat and mass transfer processes through chemical reactions.
18
CHAPTER-4
SIMULATION BY THERMODYNAMIC MODELLING
Figure 4-1: Process flow diagram of MDEA system as developed in ASPEN Plus[35]
Since the main reason for this type of modeling was to study the vapor-liquid equilibrium of
system, the open-loop simulation was modeled for easier convergence and multiple runs can be
done quickly. In order to make the simulation converged in the closed form, the various types of
design specifications were used.
In the thermodynamic model, the absorber and desorber column used are designed with specific
number of stages, where each stage is in equilibrium. In the column, the reactions taking place
are also assumed to be in equilibrium. Thus, the most optimistic results are obtained from this
simulation. In real case scenarios, the reactions are not in equilibrium and mass and heat transfer
considerations inside the column would invalidate the thermodynamic modeling. The main
reason for these simulations is for obtaining a lower limit upon the heat duty and good
estimations were provided for converging the system with the complex rate-based model.
The flow sheet shown above is developed by using ASPEN Plus software. The RADFrac
columns were selected for modeling the absorber, desorber and direct contact cooling tower
(DCC). The desorber used here has a reboiler of kettle type. A high pressure separator, known as
flash is used for separating the CO2 from water. A dot splitter was used to remove the excess
19
water and a fraction was recycled back to the desorber. The CO2 so obtained was sent to the
reciprocating compressor with 4 stages where cooling is done between stages. In each stage a
gradual compression takes place and finally CO2 was obtained at a pressure of 90 atm. The
compressor was a MCOMP model and water was decanted out in each stage. A set of pumps and
heaters were selected and used for varying pressure and temperature to the required value.
4.1 PROCEDURE FOR DEVELOPING THE FLOWSHEET
1) After providing the project title and name of the account, first the different components
were selected as shown below.
Figure 4-2: Different components selected in ASPEN PLUS
2) After selecting the components we set the base method as E-NRTL and all the properties
were defined according to it.
3) The different input streams were next given. The Flue gas was initially fed to the DCC at
a temperature and pressure of 115oC and 1 atm respectively along with cooling water at
25oC and pressure of 1 atm. MDEA was then fed at 40
oC with a pressure of 2 atm and
with a flow rate of 65 kg/hr with initial CO2 loading of 0.3 to the absorber for absorption.
Make up water was provided to the absorber at temperature of around 35oC and 2 atm
pressure to prevent entrainment amine stream.
4) After the streams were successful specified, next the input for different blocks were
provided. The input specification for the different pumps and heaters are given below:-
20
Table 4-1: Input specification of pumps
PUMPS DISCHARGE PRESSURE(atm)
PUMPRICH 3
REFPUMP 3.5
CO2PUMP 130
Table 4-2: Input specification of heaters
HEATERS TEMPERATURE(℃) PRESSURE(atm)
HEXRICH 50 3
HEXLEAN 50 1.1
HEXLEAN2 40 1
COOLER 30 1.5
CO2COOL 35 100
5) For the DCC, number stages were fixed at 4 with no condenser and the stage1 pressure
was set at 0.9 atm .A reflux ratio of 2 was maintained for the column. For the Absorber,
the number of stages was fixed at 12 with no condenser or reboiler and the stage1
pressure being 1.5 atm. In the case of the Desorber, 8 stages were taken with a kettle
reboiler and the stage1 pressure was set at 1.8 atm. Reboiler duty was taken as
5655KJ/Kg. For the desorber a pumparound was defined. It represented an internal reflux
from a hotter stage to a cooler one. In this case the pumparound was given from 8th
stage
to the 1st stage.
6) A blower was also used to raise the pressure of the cooled flue gas. It was modeled as an
isentropic blower with discharge pressure of 2 atm.
7) A 4 stage reciprocating compressor was selected and modeled as isentropic one. The
outlet stream was obtained at 90atm and 50oC. From all the other stages water was
purged out.
8) The required reactions mentioned earlier were given and the temperature dependent
constants were provided according to the table given below:-
21
Table 4-3: Reaction data used for the reactions [35]
REACTION A B C D
2.2 132.89 -13445.9 -22.47 0
2.3 231.46 -12092.1 -36.78 0
2.4 216.05 -12431.7 -35.48 0
2.5 24.76 -8989 0 0
9) After providing all the input information the different design specification were given for
optimizing and converging the simulations.
The design specifications used in the flowsheet are given below [35]:
1. Vent CO2: The amount of CO2 that is coming out from the top of the absorber is
regulated by a design specification and the desired extent of capture can be achieved by
varying the amine flow rate to the absorber. The above design specification can be
specified by selecting the Flowsheeting Options tab.
For this, first the mass flow rate of AMINEIN and VENT stream were defined. The mass
flow rate of the VENT stream was desired to be 1.2 kg/hr assuming 90 % absorption
takes place in the absorber with a tolerance of 0.05. The AMINEIN flow rate was varied
from 50-70 kg/hr to achieve this target.
2. Reboiler Duty: The reboiler duty provides the necessary heat required to produce steam
which carries the CO2 upwards from the rich amine and helps in stripping .Since the
flowsheet is an open loop one, we need to match the composition of the stream coming
out of the second heat exchanger after desorber, COOLLEAN with AMINEIN stream for
ensuring the convergence of the simulation in closed loop manner. This is achieved by
matching the loadings of the two streams. For matching the loading we need to first
define two property sets FAPPCO2 and FAPPMDEA in Prop sets tab under the Property
option. These property sets are used for calculating the apparent molar flow rate of CO2
and MDEA in the stream. The ratio of FAPPCO2 to FAPPMDEA is the loading of the
stream. The property was set to the desired value (the value in the AMINEIN stream) and
the reboiler duty was varied until the desired value is achieved.
After the prop sets have been defined, design specification of the desorber was specified.
This design specification checks the property ratio or the ratio of the two prop sets of the
LEAN stream coming out of the desorber. The reboiler duty was varied from 3000 to
6000 KJ/Kg to reach the desired loading of 0.3 as of the AMINEIN stream.
3. Cross-exchanger heat duty: In the flowsheet shown above, the two separate heat
22
exchangers – HEXRICH and HEXLEAN function together as a cross heat exchanger for
ensuring that the closed loop is not formed. However, the heat duties of both the heat
exchangers are matched so that they can effectively function as the cross-heat exchanger.
This was achieved by using the above design specification in the Flowsheeting Options
tab. The heat duty of the HEXLEAN was set equal to the negative of the heat duty of the
HEXRICH.
Initially the heat duties of the two heat exchangers were defined. Since we have assumed
that the heat duties of HEXRICH and HEXLEAN are equal but of opposite sign it was
specified that the sum of the heat duties is desired to be 0 with a tolerance of 0.02. The
mass flow rate of AMINEIN stream was varied from 50-70 kg/hr to satisfy the desired
condition.
4. Cross-heat exchanger approach temperature: Along with matching the heat duties of
heat exchangers HEXLEAN and HEXRICH, it was necessary to specify a design
specification on the temperature approach of the cross-exchanger. For this purpose, the
temperature of the inlet stream of HEXRICH is set to 10°C below the temperature of the
outlet stream of HEXLEAN.
For this purpose first the temperature of LEANCOOL and RICHPUMP were defined. It
was specified that the temperature of LEANCOOL to be 10oC more than the temperature
of RICHPUMP with a tolerance of 0.1. To achieve this we need to vary the AMINEIN
mass flow rate from 50-70 kg/hr.
5. Water balance: Water balance in the plant has to be maintained in order to have
convergence in closed loop manner. Therefore, the flow rate of the MKWATER stream
was varied to achieve the overall water balance.
To ensure water balance we first need to define the mass flow rate of CW, CW1,
MKWATER, 1, 2 and 3 stream was defined. It was specified that sum of the expression
(CW+MKWATER-CW1-1-2-3) should be equal to zero with a tolerance of 0.02.To attain
water balance we need to vary the MKWATER from 40-60 kg/hr.
23
CHAPTER-5
RESULTS AND DISCUSSION
After the initial run we found the CO2 conversion to be 75 %.Next we need to carry out
sensitivity analysis for the complete process to see the effect of different parameters on loading
and energy required in the form of Reboiler Duty. The results of this analysis are stated below:-
5.1 VARIATION OF L/G IN ABSORBER WITH LEAN LOADING FOR 80 %
CAPTURE
Table 5-1: Lean loading vs L/G data
Lean loading(mole of
CO2/mole of MDEA)
L/G (total molar flow of
liquid/total molar flow gas)
0.08
0.12
0.14
0.16
0.18
0.2
0.22
0.24
0.26
0.28
0.30
0.7
0.8
0.85
0.87
0.9
0.95
1
1.1
1.2
1.4
1.5
Figure 5-1: Variation of L/G with lean loading for 80% CO2 capture
0.6
0.8
1
1.2
1.4
1.6
0.05 0.1 0.15 0.2 0.25 0.3 0.35
L/G
(to
tal
mo
lar
flo
wo
f li
qu
id/t
ota
l
mo
lar
flo
wo
f g
as)
Lean loading(mole of CO2/mole of MDEA)
Variation of lean loading with L/G
24
DISCUSSION
It was observed from the above figure that with the increase in loading of the lean amine stream,
the flow rate of the amine increases. It is due to the reason that the loading is directly
proportional to its capability to absorb CO2 and with increase in loading the capacity to absorb
CO2 increases. It was also found that there is a sharp increase in the L/G or Amine flow rate
required to achieve the required amount of capture after a certain loading.
5.2 VARIATION OF REBOILER DUTY WITH L/G IN ABSORBER FOR 80 %
CAPTURE
Table 5-2: Data representing variation of reboiler duty with L/G in absorber
L/G(total molar flow of
liquid/total molar flow of gas)
Reboiler duty(kJ/kg)
0.75
0.78
0.8
0.85
0.9
0.95
1.05
1.1
1.2
1.4
1.45
5500
5000
4800
4400
3560
3580
3600
3615
3630
3640
3645
Figure 5-2: Variation of reboiler duty with L/G for 80% CO2 capture
3000
3500
4000
4500
5000
5500
6000
0.6 0.8 1 1.2 1.4 1.6
Reb
oil
er d
uty
(KJ
/Kg
)
L/G(total molar flow of liquid/total molar flow of gas)
Variation of Reboiler duty with L/G
25
DISCUSSION
It can be seen from the figure mentioned above that the reboiler duty is both directly and
indirectly proportional to Lean loading. Reboiler Duty was high at lower lean loadings, as the
equilibrium partial pressure of CO2 is very low at low loading. Therefore, for stripping the
solution to the desired lean loading in the desorber, large amount of steam is needed. Thus,
energy lost because of water condensation as well as vaporisation is high. Though, the amount of
steam that is required for vaporization decreases as the equilibrium partial pressure of CO2
increases as the lean loading increases. This results in lowering of reboiler duty. However, after
certain lean loading, the sensible heat component dominates as the flow rate of the liquid for
achieving 80% capture is increased as shown in Figure 5-2. Thus, scope for optimization of the
loading range exists to operate in.
5.3 EFFECT OF CAPTURE PERCENTAGE
Table 5-3: Variation of reboiler duty with lean loading at different capture
percentage
Lean
loading(mole
of CO2/mole of
MDEA)
Reboiler duty at 65%
CO2 capture
Reboiler duty at 85%
CO2 capture
Reboiler duty at 90%
CO2 capture
0.08
0.12
0.14
0.16
0.18
0.2
0.22
0.24
0.26
0.28
0.3
5680
4779
4380
4130
3530
3560
3570
3585
3595
3610
3620
5700
4800
4400
4150
3560
3580
3590
3615
3630
3640
3650
5730
4820
4425
4180
3600
3615
3620
3650
3670
3785
3820
26
Fig 5-3: Variation of reboiler Duty with lean loading for different capture percentage
DISCUSSION
The above figure tells us about how the capture% affects the reboiler duty. Capture extents for
65%, 85% and 90% are studied. The reboiler duty for different loadings at various capture
percentage was calculated. It was found out from the figure that the reboiler duty/kg of CO2
captured increases with the capture % . This may be due to the fact that with the increase in the
required solvent flowrate, the capture percentage increases and as a result sensible heat required
for the system also increases.
5.4 VARIATION IN REBOILER DUTY WITH DESORBER TEMPERATURE AND
DESORBER PRESSURE FOR 80 % CO2 CAPTURE
Table 5-4: Data representing variation of reboiler duty with temperature and pressure of
the desorber for 80% capture
Desorber pressure(kPa) Reboiler Duty(KJ/Kg) Desorber Temperature(K)
75
100
125
150
175
200
225
250
275
300
5700
4800
4400
4150
3800
3750
3720
3700
3660
3620
370
378
383
387
391
394
396
399
402
405
3400
3900
4400
4900
5400
5900
0.06 0.16 0.26
Reb
oil
er d
uty
(KJ
/Kg
)
Lean loading(mole of CO2/mole of MDEA)
Variation of Reboiler duty with lean loading
Reboiler duty at 65%
CO2 capture
Reboiler duty at 85%
CO2 capture
Reboiler duty at 90%
CO2 capture
27
Fig 5-4: Graph showing variation of reboiler duty with desorber pressure
Fig 5-5: Graph showing variation of desorber temperature with desorber pressure
DISCUSSION
The process performance is very sensitive to desorber temperature and pressure.The pressure is
directly related to temperature. With the increase in temperature pressure also increases. As at
high temperature and pressure transfer of CO2 takes place, it is desired to maintain the desorber
at high temperatures and pressures. However, it should be kept in mind that with increase in
3000
3500
4000
4500
5000
5500
6000
50 100 150 200 250 300 350
Reb
oil
er D
uty
(KJ
/Kg
)
Desorber pressure(kPa)
Variation of Reboiler Duty with Desorber Pressure
365
370
375
380
385
390
395
400
405
410
50 100 150 200 250 300 350
Des
orb
er T
emp
era
ture
(oC
)
Desorber pressure(kPa)
Variation of Desorber Temperature with Desorber
Pressure
28
temperature degradation of solvent (MDEA) takes place. Therefore, it is recommended to
maintain the desorber within the temperatrure of 130°C.
5.5 VARIATION IN REBOILER DUTY WITH LEAN LOADING FOR APPROACH
TEMPERATURES OF 5 &10 O
C
Table 5-5: Data representing variation of reboiler duty with lean loading for different
temperature approach
Lean Loading(mol of
CO2/mole of MDEA)
Reboiler Duty For 10 O
C
Cross Heat
Approach(KJ/Kg)
Reboiler Duty For 5 O
C
Cross Heat
Approach(KJ/Kg)
0.08
0.12
0.14
0.16
0.18
0.2
0.22
0.24
0.26
0.28
0.3
5700
4800
4400
4150
3560
3580
3590
3615
3630
3640
3650
5700
4800
4400
4150
3500
3480
3450
3470
3490
3510
3530
Fig 5-6: Graph showing variation of reboiler duty with cross-heat exchanger approach
3000
3500
4000
4500
5000
5500
6000
0.06 0.16 0.26 0.36
Reb
oil
er d
uty
(KJ
/Kg
)
Lean loading (mole of CO2/mol of MDEA)
Variation of reboiler duty with lean loading for 5oC and
10oC approach
Reboiler Duty For 10 OC
Cross Heat
Approach(KJ/Kg)
Reboiler Duty For 5 OC
Cross Heat
Approach(KJ/Kg)
29
DISCUSSION
The effectiveness of cross-heat exchangers was found out from temperature approach. Generally
the cold side approach is more than hot side. From the above figure it was found that Reboiler
duty is less for 5 O
C approach as compared to 10 O
C approach. Reboiler duty at higher lean
loadings for 5 O
C is almost constant. We are not getting variation in reboiler duty. It is not that
much effective. Therefore, 10 O
C approach is used.
30
CHAPTER-6
6.1 CONCLUSION
To carry out simulation and modeling for the CO2 capture process with MDEA, the E-NRTL
model was used to correlate vapor-liquid equilibrium properties of CO2-H2O-MDEA system
with the experimental data available for the CO2-H2O-MDEA system. The model is validated to
predict vapor-liquid equilibrium (VLE), Reboiler duty, and partial pressure of CO2 of the
MDEA-H2O-CO2 system with temperature ranging from 40OC to 120
OC, concentration of
MDEA up to 30wt %, and loadings of CO2 close to 1. The model provides a representation for
thermodynamic property for the CO2- H2O-MDEA system over a wider range of conditions and
gives more-reliable predictions than those from previous works.
The models were confirmed to be able to represent thermodynamic and thermal properties over a
broad range of conditions. Overall the results presented above indicate that the developed model,
through simultaneous regression to different kind of data, gave a set of optimum parameters for
CO2-MDEA-H2O.The developed models adequately represent thermodynamic and thermal
properties of the systems. Modeling results show that the developed models improved
significantly over previously existing models. The relation of CO2 capture on the lean loading,
desorber temperature and pressure, and reboiler duty, was sufficient to ensure the completion of
the simulation and designing of the Carbon Capture and Storage project work.
The system was well represented by the E-NRTL model. However, because of high complexity
of model equations, the implementation becomes difficult.
6.2 FUTURE SCOPE OF WORK
A recurring theme in this work is that there are no rules of thumb which always apply; every
situation is different. Many rules of thumb are presented in the literature and most of them are
based on established criterion and work. However, just because they work in most situations
does not mean they are always the best choice. This work should be considered as a starting
point, a first guess. A thorough evaluation should be performed to optimize performance. The
benefits of deviating from these guidelines should be weighed against the risks.
As the rate of reaction of CO2 with MDEA is low, the thermodynamic model should be further
developed for aqueous MDEA blended with PZ or other additives like Sulfolane.
Accurate simulation can provide increased ability to anticipate the effect of changing treating
conditions in the plant.The future work should emphasize on high percent removal of the gas for
optimum conditions.
31
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